Multiple-stage catalyst system for self-metathesis with controlled isomerization and cracking

ABSTRACT

Embodiments of processes and multiple-stage catalyst systems for producing propylene comprising introducing a hydrocarbon stream comprising 2-butene to an isomerization catalyst zone to isomerize the 2-butene to 1-butene, passing the 2-butene and 1-butene to a metathesis catalyst zone to cross-metathesize the 2-butene and 1-butene into a metathesis product stream comprising propylene and C 4 -C 6  olefins, and cracking the metathesis product stream in a catalyst cracking zone to produce propylene. The isomerization catalyst zone comprises a silica-alumina catalyst with a ratio by weight of alumina to silica from 1:99 to 20:80. The metathesis catalyst comprises a mesoporous silica catalyst support impregnated with metal oxide. The catalyst cracking zone comprises a mordenite framework inverted (MFI) structured silica catalyst.

CROSS-REFERENCE TO RELATED APPLICATIONS

This application claims the benefit of U.S. Provisional Application Ser.No. 62/448,478 filed Jan. 20, 2017, incorporated herein by reference.

TECHNICAL FIELD

Embodiments of the present disclosure generally relate to propyleneproduction, and more specifically relate to converting butene topropylene using a multiple-stage catalyst system comprisingisomerization, metathesis, and cracking catalysts.

BACKGROUND

In recent years, there has been a dramatic increase in the demand forpropylene to feed the growing markets for polypropylene, propylene oxideand acrylic acid. Currently, most of the propylene produced worldwide(74 million tons/year) is a by-product from steam cracking units (57%)which primarily produce ethylene, or a by-product from Fluid CatalyticCracking (FCC) units (30%) which primarily produce gasoline. Theseprocesses cannot respond adequately to a rapid increase in propylenedemand.

Other propylene production processes contribute about 12% of totalpropylene production. Among these processes are propane dehydrogenation(PDH), metathesis reactions requiring both ethylene and butene, highseverity FCC, olefins cracking and methanol to olefins (MTO). However,propylene demand has exceeded ethylene and gasoline/distillate demand,and propylene supply has not kept pace with this increase in propylenedemand.

SUMMARY

Accordingly, ongoing needs exist for improved processes for theselective production of propylene using multiple-stage catalyst systems.Embodiments of the present disclosure are directed to propyleneproduction from butenes by a multiple-stage catalyst system.

In one embodiment, a process for the production of propylene isprovided. The process comprises introducing a hydrocarbon streamcomprising 2-butene to an isomerization catalyst zone to isomerize the2-butene to 1-butene, where the isomerization catalyst zone comprises asilica-alumina catalyst with a ratio by weight of alumina to silica from1:99 to 20:80. The process also includes passing the 2-butene and1-butene to a metathesis catalyst zone to cross-metathesize the 2-buteneand 1-butene into a metathesis product stream comprising propylene,unconverted C₄, and higher metathesis product C₅ and C₆ olefins, wherethe metathesis catalyst comprises a mesoporous silica catalyst supportimpregnated with metal oxide. The process further includes cracking themetathesis product stream in a catalyst cracking zone to producepropylene, where the catalyst cracking zone comprises a mordeniteframework inverted (MFI) structured silica catalyst.

In another embodiment, a multiple-stage catalyst system for producingpropylene from from a hydrocarbon stream comprising 2-butene isprovided. The multiple-stage catalyst system comprises an isomerizationcatalyst zone, a metathesis catalyst zone downstream of theisomerization zone, and a cracking catalyst zone downstream of themetathesis catalyst zone. The isomerization catalyst zone comprises asilica-alumina catalyst with a ratio by weight of alumina to silica from1:99 to 20:80, where the silica-alumina catalyst zone isomerizes the2-butene to 1-butene. The metathesis catalyst zone comprises amesoporous silica catalyst support impregnated with metal oxide to forma mesoporous silica catalyst, where the mesoporous silica catalyst zonecross-metathesizes the 2-butene and 1-butene into a metathesis productstream comprising propylene, unconverted C₄, and higher metathesisproduct C₅ and C₆. The cracking catalyst zone comprises a mordeniteframework inverted (MFI) structured silica catalyst, where the crackingcatalyst zone cracks the metathesis product stream to produce propylene.

Additional features and advantages of the described embodiments will beset forth in the detailed description which follows, and in part will bereadily apparent to those skilled in the art from that description orrecognized by practicing the described embodiments, including thedetailed description which follows, the claims, as well as the appendeddrawings.

BRIEF DESCRIPTION OF THE DRAWINGS

FIG. 1 is an X-ray Powder Diffraction (XRD) graph illustrating the XRDprofile of silica-alumina catalysts with varying silica to aluminaratios, in accordance with one or more embodiments of the presentdisclosure.

FIG. 2 is an XRD graph of tungsten and a mesoporous silica catalystcomprising a silica support impregnated with tungsten, in accordancewith one or more embodiments of the present disclosure.

FIG. 3 is a graph illustrating the performance of a multiple-stagecatalyst system over time and reaction temperature changes, inaccordance with one of more embodiments of the present disclosure.

DETAILED DESCRIPTION

Embodiments of the present disclosure are directed to systems andmethods for converting a hydrocarbon stream comprising 2-butene to astream comprising propylene by catalyzed butene isomerization, catalyzedmetathesis, and catalyzed cracking. Specifically, the presentembodiments are related to a multiple-stage (for example, three-stage)catalyst system containing isomerization, metathesis, and crackingcatalysts for greater propylene (C₃═) production from a butene stream.While a three-stage catalyst system with 3 catalysts is used throughoutthis disclosure for simplicity and clarity, it may be appreciated thatthe multiple-stage catalyst system may include more than 3 catalystsincluding 4 catalysts, 5 catalysts, or 6 or more catalysts. In one ormore embodiments, the isomerization catalyst is followed by themetathesis catalyst, and the metathesis catalyst is followed by thecracking catalyst to provide a greater yield of propylene, andoptionally a greater combined yield of propylene and ethylene.

The hydrocarbon stream may be any stream comprising 2-butene. Forexample, the hydrocarbon stream may be a raffinate stream created from anaphtha cracking process or an FCC stream. Such a stream may be aRaffinate 1 stream, a Raffinate 2 stream or a Raffinate 3 stream.Raffinate 1 is the residual stream which is obtained when a C₄ streamfrom a naphtha cracking process or from a gas cracking process, forexample, is subjected to the removal of 1,3-butadiene therefrom. The C₄stream typically contains, as its chief components, n-butene, 1-butene,2-butene, isobutene and 1,3-butadiene, and optionally some isobutanewith the chief components together forming up to 99% or more of the C₄stream. Removal typically may be by extractive distillation with anaprotic solvent such as acetonitrile, N-methylpyrrolidone orN,N-dimethylformamide. Any remaining butadiene after extractivedistillation may optionally being removed by an additional treatmentsuch as selective hydrogenation. The resulting and remaining residualstream stream is a Raffinate 1 stream and contains 1-butene, 2-buteneand isobutene. A Raffinate 2 stream in turn is a mixture of 1-butene and2-butene which remains when a Raffinate 1 stream has the isobuteneseparated therefrom. The separation may be by hydrogenation totert-butanol in the presence of sulphuric acid, by reaction of theRaffinate 1 stream with methanol to synthesize methyl tert-butyl ether,or by oligomerization or polymerization of the isobutene, for example.Further, a Raffinate 3 stream is what is obtained when the 1-butene in aRaffinate 2 stream is separated therefrom. The seperation may be byfractionation, extractive distillation or molecular sieve absorption,for example. The residual stream of cis 2-butene and trans 2-butene isthe Raffinate 3 stream.

In one or more embodiments, the hydrocarbon stream is a Raffinate 2stream from a fluid catalytic cracker or an ethylene cracker. TheRaffinate 2 stream may comprise various compositions. Non-limitingexamples of Raffinate 2 stream compositions include 45 to 55 weightpercentage (wt %) 1-butene, 20 to 30 wt % 2-butene, 10 to 20 wt %n-butane, 5 to 15 wt % iso-butane, and 0 to 5 wt % other components; 10to 20 wt % 1-butene, 20 to 30 wt % 2-butene, 8 to 18 wt % n-butane, 37to 47 wt % iso-butane, and 0 to 8 wt % other components; 48 to 50 weightpercentage (wt %) 1-butene, 25 to 37 wt % 2-butene, 14 to 16 wt %n-butane, 9 to 10 wt % iso-butane, and 0 to 1 wt % other components; 15to 17 wt % 1-butene, 25 to 27 wt % 2-butene, 11 to 13 wt % n-butane, 41to 44 wt % iso-butane, and 2 to 6 wt % other components; approximately49.6 wt % 1-butene, approximately 26.0 wt % 2-butene, approximately 14.7wt % n-butane, approximately 9.4 wt % iso-butane, and approximately 0.3wt % other components; or approximately 15.6 wt % 1-butene,approximately 26.2 wt % 2-butene, approximately 12.0 wt % n-butane,approximately 42.1 wt % iso-butane, and approximately 4.1 wt % othercomponents

As used in this disclosure, a “reactor” refers to a vessel in which oneor more chemical reactions may occur between one or more reactantsoptionally in the presence of one or more catalysts. For example, areactor may include a tank or tubular reactor configured to operate as abatch reactor, a continuous stirred-tank reactor (CSTR), or a plug flowreactor. Example reactors include packed bed reactors such as fixed bedreactors, and fluidized bed reactors. One or more “reaction zones” maybe disposed in a reactor. As used in this disclosure, a “reaction zone”refers to an area where a particular reaction takes place in a reactor.For example, a packed bed reactor with multiple catalyst beds may havemultiple reaction zones, where each reaction zone is defined by the areaof each catalyst bed.

As shown in Formula 1 as follows, isomerization of 2-butenes to1-butenes forms an equilibrium as denoted by the bi-directional arrow.The isomerization is achieved with the isomerization catalyst. Thencross-metathesis is achieved as shown in Formula 2 with the metathesiscatalyst. Cross-metathesis is an organic reaction that entails theredistribution of fragments of alkenes by the scission and regenerationof carbon-carbon double bonds. In the case of 2-butene and 1-butene, theredistribution results in propylene and C₅-C₆ olefins. Propylene mayalso be formed from a secondary cross-metathesis reaction betweenethylene and 2-butene as shown in Formula 1. The metathesis catalyst canalso result in “self-metathesis” as shown in Formula 4. Without wishingto be bound by theory, it is believed 1-butene reacts with itself due tomoderate or less active sites providing room to react with itselfinstead of forming an isomer such as 2-butene. However, the tendency ofcross-metathesis occurring is much higher that self-metathesis. Further,as shown in Formula 5, “catalyzed cracking” refers to the conversion ofmainly C₅/C₆ alkenes and unconverted C₄'s from the metathesis reactionto propylene and ethylene (C₂═ and C₃═), some light gases (C₁, C₂), andalso some higher hydrocarbons depending on cracking conditions.

Referring to Formulas 1-5, the “isomerization,” “metathesis,” and“catalytic cracking” reactions are not limited to these reactants andproducts; however, Formulas 1-5 provide a basic illustration of thereaction methodology. As shown in Formulas 2-4, metathesis reactionstake place between two alkenes. The groups bonded to the carbon atoms ofthe double bond are exchanged between the molecules to produce two newalkenes with the swapped groups. The specific catalyst that is selectedfor the olefin metathesis reaction may generally determine whether acis-isomer or trans-isomer is formed, as the coordination of the olefinmolecules with the catalyst play an important role, as do the stericinfluences of the substituents on the double bond of the newly formedmolecule.

In operation, a product stream comprising propylene is produced from abutene containing stream by isomerization, metathesis conversion, andcracking by contacting the butene stream with the multiple-stagecatalyst system. The butene stream may comprise 2-butene, and optionallycomprises one or more isomers, such as 1-butene, trans-2-butene, andcis-2-butene. The present discussion centers on butene based feedstreams; however, it is known that other C₁-C₆ components may also bepresent in the feed stream.

In one or more embodiments, the present multiple-stage catalyst systemcomprises: a silica-alumina catalyst, a mesoporous silica catalystsupport impregnated with metal oxide downstream of the silica-aluminacatalyst; and a mordenite framework inverted (MFI) structured silicacatalyst downstream of the mesoporous silica catalyst. As indicatedsupra the discussion of a three-stage catalyst system with 3 catalystsis merely for simplicity and a multiple-stage catalyst system with 4 ormore catalysts is also envisioned. The silica-alumina catalyst is anisomerization catalyst which facilities isomerization of 2-butene to1-butene. The mesoporous silica catalyst, which is downstream of thesilica-alumina catalyst, is a metathesis catalyst which also facilitatesisomerization of 2-butene to 1-butene followed by cross-metathesis ofthe 2-butene and 1-butene into a metathesis product stream comprisingpropylene, and other alkenes/alkanes such as pentene. The MFI structuredsilica catalyst, which is downstream of the metathesis catalyst, is acracking catalyst which produces propylene from C₄ or C₅ olefins in themetathesis product stream, and may also yield ethylene.

In one or more embodiments, a first stage comprising the silica-aluminacatalyst assists in the isomerization between 2-butene and 1-butene. Thesecond stage comprising the mesoporous silica catalyst contains a metaloxide impregnated on a silica support. The mesoporous silica catalystperforms the butenes to propylene cross-metathesis. The mesoporoussilica catalyst may also perform self-metathesis by both isomerizing andmetathesizing the butenes. Self-metathesis with the mesoporous silicacatalyst is not mandatory for the reaction to proceed as thesilica-alumina catalyst already performs the isomerization of thebutenes for butene to propylene cross-metathesis by the mesoporoussilica catalyst. The third stage comprises the MFI structured catalystwhich converts the unreacted C₄ hydrocarbons and the produced C₅+hydrocarbons to lighter olefins such as ethylene and propylene. In oneor more embodiments, the MFI structured catalyst is a ZSM-5 crackingcatalyst.

The ratio of silica to alumina in the silica-alumina catalyst may bevaried to provide differing surface area, pore volume, and isomerizationperformance. In one or more embodiments, the silica-alumina catalyst hasa ratio by weight of alumina to silica from about 1:99 to about 20:80.In further embodiments, the silica-alumina catalyst has a ratio byweight of alumina to silica from about 1:99 to about 15:85, or about2:98 to 10:90, or about 2:98 to 8:12, or about 3:97 to 7:93. It isdesirable to minimize the alumina percentage as adding more aluminaaffects the deactivation and fouling rate of the silica-aluminacatalyst. Increased amounts of alumina contribute to deactivation of thesilica-alumina catalyst while not significantly enhancing theisomerization activity. As demonstrate by the experimental results inthe Examples section of this disclosure, whether you add 5% alumina or75% alumina to the support of the silica-alumina catalyst, theisomerization activity remains substantially steady.

Maximizing the isomerization of 2-butene to 1-butene by thesilica-alumina catalyst improves the overall yield of propylene from themultiple-stage catalyst system. An increased isomerization by thesilica-alumina catalyst ensures sufficient availability of both 2-buteneand 1-butene for the cross-metathesis reaction by the mesoporous silicacatalyst resulting in an ultimate increase in the metathesis productstream for cracking into propylene by the MFI structured silicacatalyst. The isomerization of the hydrocarbon feed in the isomerizationcatalyst zone with the silica-alumina catalyst adjusts the ratio of thevarious butene isomers (1-butene, 2-cis-butene, trans-2-butene, etc.) toan optimal ratio to perform the subsequent metathesis reactions. Theoptimal ratio of the various butene isomers directs the reaction towardcross-metathesis instead of undesirable side and skeletal reactions.When the isomerization catalyst zone of the silica-alumina catalystlayer is added, double-bond isomerization reaction occurs, convertingthe hydrocarbon feed regardless of the initial ratio of 1-butenes and2-butenes to the optimal ratio. After the isomerization catalyst zone,the metathesis reactions occur in the metathesis catalyst zone. As thefeed is already in the right ratio, the metathesis reaction readilyoccurs in the metathesis catalyst zone of mesoporous silica catalyst.The undesirable skeletal reactions are reduced because there is reducedresidual butene left for these undesired reactions to utilize. Thecross-metathesis and self-metathesis reactions are faster and morefavored than the side reactions. The cross-metathesis is also morefavored over the self-metathesis which results in more propylene yieldthan ethylene yield.

The isomerization may be completed at a broad range of temperatures bythe silica-alumina catalyst. The broad temperature range of thesilica-alumina catalyst allows the operating temperature of the entiremultiple-stage catalyst system to be decreased in comparison to a dualstage catalyst system without the silica-alumina catalyst layer whichrelies upon self-metathesis by the mesoporous silica catalyst forisomerization of the butenes. Without wishing to be bound by theory, itis believed that the silica-alumina catalyst is capable of performingthe isomerization of the butenes at a lesser temperature than theisomerization aspect of self-metathesis by the mesoporous silicacatalyst. The lesser reactor operating temperature allowed by thesilica-alumina catalyst for isomerization results in an increase inultimate propylene yield from the multiple-stage catalyst system withoutcreating undesired products, such as isobutene.

Undesirable products may form at higher temperatures as the catalyst inthe isomerization zone contains only silica and alumina. Alumina isknown to deactivate and crack the feed at higher temperatures,especially when the percent of alumina content is high. When thetemperature is increased, the alumina starts to crack the butenes topropylene and ethylene which is undesirable. Also, small amounts of C5sand C6s may be formed in the product. The presence of the undesirableproducts in the reactor in the metathesis catalyst zone reduces thepropylene as well as ethylene rate of formation. Metathesis reactionsare controlled by equilibrium, so avoiding propylene, pentene andhexenes in considerable amounts prior to the metathesis catalyst resultsin preferred propylene formation. Propylene, pentene and hexenes inconsiderable amounts prior to the metathesis catalyst results in areaction direction shift from producing primarily propylene toproduction of butenes from the propylene, pentenes and hexanes present.Furthermore, skeletal isomerization reactions are temperature sensitivewhere the higher the temperatures, the higher the skeletal activitywhich results in increased isobutene. When present, isobutene activatesside reactions producing undesired products. Besides the undesiredproducts listed previously, running the isomerization catalyst zone at alower temperature correlates to lesser operating costs due to thereduction in the required heating.

In one or more embodiments, the total pore volume of the silica-aluminacatalyst may be from about 0.600 cm³/g to about 2.5 cm³/g, or about0.600 cm³/g to about 1.5 cm³/g, or about 0.600 cm³/g to about 1.3 cm³/g,or about 0.600 cm³/g to about 1.1 cm³/g, or about 0.700 cm³/g to about1.1 cm³/g, or about 0.800 cm³/g to about 1.3 cm³/g, or about 0.900 cm³/gto about 1.2 cm³/g.

Moreover, while broader ranges are contemplated, the silica-aluminacatalyst may, in one or more embodiments, include a surface area ofabout 200 meters²/gram (m²/g) to about 600 m²/g. In further embodiments,the silica-alumina catalyst may have a surface area of from about 225m²/g to about 350 m²/g, or about 225 m²/g to about 325 m²/g, or about250 m²/g to about 325 m²/g.

The acidity of the silica-alumina catalyst may be controlled with theamount and selection of alumina precursor in the catalyst precursorsolution. Isomerization is affected by the acidity of the silica-aluminacatalyst. Acidity is controlled in at least two ways. First, the totalnumber of acidic sites in the silica-alumina catalyst is controlled bythe amount of aluminum incorporated into the structure. The morealuminum sites that are present the more Al—OH will be present. Second,the acid strength is also affected by the aluminum sites and how theyinteract with the silica sites. The source of alumina in thesilica-alumina catalyst may have an effect on the formation of varioussites. For example: fumed alumina has a large cluster of alumina alreadyformed, therefore the interactions between the alumina and silica arelargely predefined and limited to the interface of the two discretematerials. In the case of Al(NO₃)₃ the alumina that is created is asingle molecule and can potentially interact with silica in alldimensions remaining isolated. Further, in various embodiments, thesilica-alumina catalyst may have a total acidity of up to about 0.5millimole/gram (mmol/g), or about 0.01 mmol/g to about 0.5 mmol/g, orabout 0.1 mmol/g to about 0.5 mmol/g, or about 0.3 mmol/g to about 0.5mmol/g, or about 0.4 mmol/g to about 0.5 mmol/g. It will be appreciatedthat in further embodiments the silica-alumina catalyst may have a totalacidity below 0.01 mmol/g or above 0.5 mmol/g.

Various structures are contemplated for the mesoporous silica catalystsupport, for example, a molecular sieve or a zeolite. As used in theapplication, “mesoporous” means that the silica catalyst support has anarrow pore size distribution. Specifically, the mesoporous silicacatalyst support includes a narrow pore size distribution of from about2.5 nm to about 40 nm and a total pore volume of at least about 0.60cm³/g. Without being bound by theory, the present pore size distributionand pore volume are sized to achieve better catalytic activity andreduced blocking of pores by metal oxides, whereas smaller pore volumeand pore size catalyst systems are susceptible to pore blocking andthereby reduced catalytic activity.

In one or more embodiments, the pore size distribution of the mesoporoussilica catalyst support may range from about 2.5 nm to about 40 nm, orabout 2.5 nm to about 20 nm, or about 2.5 nm to about 4.5 nm, or about2.5 nm to about 3.5 nm, or about 8 nm to about 18 nm, or about 12 nm toabout 18 nm.

In one or more embodiments, the total pore volume of the mesoporoussilica catalyst may be from about 0.600 cm³/g to about 2.5 cm³/g, orabout 0.600 cm³/g to about 1.5 cm³/g, or about 0.600 cm³/g to about 1.3cm³/g, or about 0.600 cm³/g to about 0.900 cm³/g, or about 0.700 cm³/gto about 0.900 cm³/g, or about 0.800 cm³/g to about 1.3 cm³/g.

Moreover, while broader ranges are contemplated, the mesoporous silicacatalyst may, in one or more embodiments, include a surface area ofabout 200 m²/g to about 600 m²/g. In further embodiments, the mesoporoussilica catalyst may have a surface area of from about 225 m²/g to about350 m²/g, or about 225 m²/g to about 325 m²/g, or about 250 m²/g toabout 325 m²/g, or about 250 m²/g to about 300 m²/g.

Further, the mesoporous silica catalyst may have a total acidity of upto about 0.5 millimole/gram (mmol/g), or about 0.01 mmol/g to about 0.5mmol/g, or about 0.1 mmol/g to about 0.5 mmol/g, or about 0.3 mmol/g toabout 0.5 mmol/g, or about 0.4 mmol/g to about 0.5 mmol/g. Acidity isgenerally maintained at or less than about 0.5 mmol/g to yield thedesired selectivity of propylene and reduced production of undesirablebyproducts such as aromatics. Increasing acidity may increase theoverall butene conversion; however, this increased conversion may leadto less selectivity and increased production of aromatic byproducts,which can lead to catalyst coking and deactivation.

Furthermore, the mesoporous silica catalyst may have a particle size offrom about 20 nm to about 200 nm, or about 50 nm to about 150 nm, orabout 75 nm to about 125 nm. In additional embodiments, the mesoporoussilica catalyst may have an individual crystal size of about 1 μm toabout 200 μm, or about 10 μm to about 150 μm or about 50 μm to about 120μm.

The catalyst of the cross-metathesis reaction is the impregnated metaloxide of the mesoporous silica catalyst. The metal oxide may compriseone or more oxides of a metal from the Groups 6-10 of the IUPAC PeriodicTable. In one or more embodiments, the metal oxide may be an oxide ofmolybdenum, rhenium, tungsten, or combinations thereof. In a specificembodiment, the metal oxide is tungsten oxide (WO₃). It is contemplatedthat various amounts of metal oxide may be impregnated into themesoporous silica catalyst support. For example and not by way oflimitation, the weight percentage (wt. %) of metal oxide, for example,WO₃, in the mesoporous silica catalyst is about 1 to about 30 wt. %, orabout 1 to about 25 wt. %, or about 5 to about 20 wt. %, or about 5 toabout 15 wt. %, or about 8 to about 12 wt. %. The weight percentage oftungsten in the silica support may be measured by X-ray fluorescence(XRF) or Inductively coupled plasma (ICP).

Additionally, various silica structures are contemplated for the MFIstructured silica catalyst. For example, the MFI structured silicacatalyst may include MFI structured aluminosilicate zeolite catalysts orMFI structured silica catalysts free of alumina. As used in thisdisclosure, “free” means less than 0.001% by weight of alumina in theMFI structured silica catalyst. Moreover, it is contemplated that theMFI structured silica catalyst may include other impregnated metaloxides in addition to or as an alternative to alumina. Like themesoporous silica catalyst, the MFI structured catalysts may havealumina, metal oxides, or both impregnated in the silica support. Inaddition to or as a substitute for alumina, it is contemplated toinclude the metal oxides listed prior, specifically, one or more oxidesof a metal from Groups 6-10 of the IUPAC Periodic Table, morespecifically, metal oxides of molybdenum, rhenium, tungsten, titanium,or combinations thereof.

For the MFI structured aluminosilicate zeolite catalysts, variousamounts of alumina are contemplated. In one or more embodiments, the MFIstructured aluminosilicate zeolite catalysts may have a molar ratio ofsilica to alumina of about 5 to about 5000, or about 100 to about 4000,or about 200 to about 3000, or about 1000 to about 2500, or about 1500to about 2500. Various suitable commercial embodiments of the MFIstructured aluminosilicate zeolite catalysts are contemplated, forexample, ZSM-5 zeolites such as MFI-280 produced by ZeolystInternational or MFI-2000 produced by Saudi Aramco.

Various suitable commercial embodiments are also contemplated for thealumina free MFI structured catalysts. One such example is Silicalite-1produced by Saudi Aramco.

The MFI structured silica catalyst may include a pore size distributionof from about 1.5 nm to 3 nm, or about 1.5 nm to 2.5 nm. Furthermore,the MFI structured silica catalyst may have a surface area of from about300 m²/g to about 425 m²/g, or about 340 m²/g to about 410 m²/g.Additionally, the MFI structured silica catalyst may have a totalacidity of from about 0.001 mmol/g to about 0.1 mmol/g, or about 0.01mmol/g to about 0.08 mmol/g. The acidity is maintained equal to or lessthan about 0.1 mmol/g in order to reduce production of undesirablebyproducts such as aromatics. Increasing acidity may increase the amountof cracking; however, this increased cracking may also lead to lessselectivity and increased production of aromatic byproducts, which canlead to catalyst coking and deactivation.

In some cases, MFI structured silica catalyst may be modified with anacidity modifier to adjust the level of acidity in the MFI structuredsilica catalyst. For example, these acidity modifiers may include rareearth modifiers, phosphorus modifiers, potassium modifiers, orcombinations thereof. However, as the present embodiments are focused onreducing the acidity to a level at or below 0.1 mmol/g, the presentstructured silica catalyst may be free of acidity modifier, such asthose selected from rare earth modifiers, phosphorus modifiers,potassium modifiers, or combinations thereof. As used in thisdisclosure, “free of acidity modifiers” means less than less than 0.001%by weight of acidity modifier in the MFI structured silica catalyst.

Further, the MFI structured silica catalyst may have a pore volume offrom about 0.1 cm³/g to about 0.3 cm³/g, or about 0.15 cm³/g to about0.25 cm³/g. Additionally, the MFI structured silica catalyst may have anindividual crystal size ranging from about 10 nm to about 40 μm, or fromabout 15 μm to about 40 μm, or from about 20 μm to about 30 μm. Inanother embodiment, the MFI structured silica catalyst may have anindividual crystal size in a range of from about 1 μm to about 5 μm.

Moreover, various amounts of each catalyst are contemplated for thepresent multiple-stage catalyst system. For example, it is contemplatedthat the ratio by volume of the silica-alumina catalyst, the mesoporoussilica catalyst, and the MFI structured silica catalyst may range fromabout 5:1:1 to about 1:5:1 to about 1:1:5, or about 2:1:1 to about 1:2:1to about 1:1:2, or about 1:1:1.

It is contemplated that the isomerization, the metathesis catalyst, andthe cracking catalyst are disposed in one reactor or multiple reactors.For example, it may be desirable to use separate reactors for one ormore of the silica-alumina catalyst, the mesoporous silica catalyst, orthe MFI structured silica catalyst when they operate at differentenvironmental conditions, including temperature and pressure. Regardlessof whether one or multiple reactors contain the multiple catalysts, themultiple-stage catalyst system will have an isomerization catalyst zoneor section, a metathesis catalyst zone or section downstream of theisomerization zone or section, and a cracking catalyst zone or sectiondownstream of the metathesis zone or section. For example, thesilica-alumina catalyst may be located in the top part of the reactor,the mesoporous silica metathesis catalyst may be located in the middlepart of the reactor, and the MFI structured silica cracking catalyst maybe disposed in the bottom part of the reactor, assuming the reactantstream enters the top portion of the reactor. For example, each catalystmay be positioned as discrete catalyst beds. Moreover, it iscontemplated that the multiple catalysts of the multiple-stage catalystsystem may be in contact with one or more of the other catalysts orseparated. However, if the multiple catalysts are in contact, it isdesirable that the isomerization catalyst is still disposed upstream ofthe metathesis catalyst and that the metathesis catalyst is stilldisposed upstream of the cracking catalyst. The catalysts can be used inthe same reactor or with different reactors arranged in series.Alternatively, it is contemplated that the isomerization catalyst(silica-alumina catalyst) is disposed in a first reactor, the metathesiscatalyst (mesoporous silica catalyst) is disposed in a separate secondreactor downstream of the first reactor, and the cracking catalyst (MFIstructured silica catalyst) is disposed in a separate third reactordownstream of the second reactor. Additionally, it is contemplated thatthe isomerization catalyst (silica-alumina catalyst) is disposed in afirst reactor, the cracking catalyst (MFI structured silica catalyst) isdisposed in a separate second reactor downstream of the first reactor,and the metathesis catalyst (mesoporous silica catalyst) is disposed inthe first reactor downstream of the isomerization catalyst or the secondreactor upstream of the cracking catalyst. In specific embodiments,there is a direct conduit between the first reactor and second reactorand the second and third reactor, so that the cracking catalyst candirectly crack the product of the butene cross-metathesis reaction.

Various methods of making the catalysts used in the multiple-stagecatalyst system are contemplated. Specifically, the processes of wetimpregnation and hydrothermal synthesis may be utilized; however, othercatalyst synthesis techniques are also contemplated.

Various operating conditions are contemplated for the contacting of thebutene stream with the multiple-stage catalyst system. For example, thebutene stream may contact the multiple-stage catalyst system at a spacehour velocity of about 10 to about 10,000 h⁻¹, or about 300 to about1200 h⁻¹. Moreover, the butene stream may contact the catalyst system ata temperature of from about 200 to about 600° C., or about 300 to about600° C. Furthermore, the butene stream may contact the catalyst systemat a pressure from about 1 to about 30 bars, or about 1 to about 10bars.

Optionally, each of the catalysts in the multiple-stage catalyst systemmay be pretreated prior to the isomerization, metathesis, and cracking.For example, the multiple-stage catalyst system may be pretreated withN₂ for about 1 hour to about 5 hours before isomerization, metathesis,and cracking at a temperature of at least about 400° C., or at leastabout 500° C.

The product stream yielded by the multiple-stage catalyst system mayhave at least an 80 mol. % conversion of butene and a propylene yield inmol. % of at least 30%. In a further embodiment, the product stream mayhave at least an 85 mol. % conversion of butene and a propylene yield inmol. % of the at least 40%. Moreover, the product stream may have atleast a 10 mol. % yield of ethylene, or at least a 15 mol. % yield ofethylene, or at least a 20 mol. % yield of ethylene. In yet anotherembodiment, the product stream may have at least 45 mol. % yield ofpropylene, or at least about a 50 mol. % yield of propylene.

Moreover, the product stream may comprise less than 1 about wt %aromatics, or less than about 5 wt % of alkanes and aromatics. Withoutbeing bound by theory, in some embodiments it may be desirable that thearomatics and alkanes yield be low as it indicates coke formation, whichmay result in catalyst deactivation.

EXAMPLES

The following examples show the preparation of various catalysts whichare used in a combination as in the present multiple catalysts.

Example 1 Preparation of SiO₂—Al₂O₃ Isomerization Catalyst

To form the SiO₂—Al₂O₃ catalyst, a total of 20 grams of Q-10 (silica)from Silysia and gamma phase aluminum oxide catalyst support (alumina)from Alfa Aesar were added to a beaker containing 50 milliliters (ml) ofdeionized (DI) water. The relative amounts of silica and alumina addedto the DI water were varied depending on the SiO₂—Al₂O₃ ratio desired.For example, in the case of 10% Al₂O₃ and 90% SiO₂, 2 grams of aluminaand 18 grams (g) of silica were added to the 50 ml of DI water. Thealumina-silica mixture was then mixed used a magnetic stirrer for 2hours (h) at 580 rotations per minute (rpm). After 2 h, the solution wasplaced in a rotary evaporator, commonly known as a rotavap. The rotavapwas rotated at 171 rpm, and operated under 292 millibar (mbar) vacuumand 80° C. Cold water at 6° C. was pumped into the rotavap housing toenhance condensation. The synthesized material was then placed in adrying oven overnight at 80° C. and calcined at 200° C. for 3 h and thenat 575° C. for 5 h. The ramping rate from 200° C. to 575° C. was 3° C.per minute (min).

The XRD patterns of the SiO₂—Al₂O₃ catalyst with varying ratios ofalumina and silica are shown in FIG. 1. The XRD indicates a broad peakcentered around 2θ=22.5° indicating the presence of silica in thematerial. Further, the XRD indicates peaks at approximately 2θ=37°, 38°,46.3°, and 67° attributable to alumina. The relative strengths of thealumina peaks and the silica peaks for a particular sample is inalignment with the relative ratio of alumina and silica in theSiO₂—Al₂O₃ catalyst.

Example 2 Preparation of Mesoporous Cross-Metathesis Catalyst (SiO₂Impregnated with a Tungsten Precursor)

In a typical synthesis, SiO₂ supports were prepared according to Example1 with 0% Al₂O₃ and 100% SiO₂. To synthesize the SiO₂ supports withimpregnated tungsten precursor, 2 g of the SiO₂ support from Example 1were placed in an 80 ml beaker. 0.235 g of ammonium metatungstatehydrate [(NH₄)6H₂W₁₂O₄₀.xH₂O] 99.99% trace metals basis was mixed with 2ml of DI water. The ammonium metatungstate hydrate was then addeddrop-wise to the 2 g of SiO₂ support. Typically, 5 drops were placed onthe SiO₂ support. A glass rod was used to thoroughly mix the support.Subsequently, the SiO₂ support mixed with the ammonium metatungstatehydrate was placed in a drying oven overnight at 80° C. The dried SiO₂support mixed with the ammonium metatungate hydrate was calcined at 250°C. for 2 h followed by calcining at 550° C. for 8 h with a ramping rateof 1° C. per min until 250° C. was reached and 3° C. per min until 550°C. was reached. This forms mesoporous silica catalyst.

The XRD patterns of tungsten oxide and a mesoporous silica catalyst(SiO₂ support with impregnated tungsten oxide) are shown in FIG. 2. Thepeaks of the XRD patterns are aligned for the tungsten oxide and themesoporous silica catalyst indicating the presence of the tungsten oxidein the mesoprous silica catalyst.

Example 3 Preparation of Silicalite-1 Cracking Catalyst

In a typical synthesis, 4.26 g tetrapropylammonium bromide (TPA) and0.7407 g ammonium fluoride was dissolved in 72 ml of water and stirredwell for 15 mins. Then, 12 g fumed silica was added and stirred welluntil homogenized. The obtained gel was autoclaved and kept at 200° C.for 2 days. The molar composition of gel was 1 SiO₂: 0.08 (TPA)Br: 0.10NH₄F: 20 H₂O. The solid products obtained were washed with water anddried at 80° C. overnight. The template was removed by calcination inair at 750° C. for 5 hours at a ramping rate of 3° C. per min.

Example 4 Preparation of MFI-2000 Cracking Catalyst

In a typical synthesis, 8.52 g TPA and 1.48 g ammonium fluoride wasdissolved in 72 ml of water and stirred well for 20 mins. 24 g fumedsilica and 0.15 g of aluminum nitrate were gradually addedsimultaneously to the TPABr solution while stirring vigorously. Once thesolution gelled, the gel was mixed vigorously for about 10 minutes untilhomogenized. The obtained gel was autoclaved and kept at 200° C. for 2days. After two days the autoclave was quenched in cold water for 30minutes. The molar composition of gel was 1 SiO₂: 0.0005 Al₂O₃: 0.08(TPA)Br: 0.10 NH₄F: 20 H₂O. The solid products obtained were washed withwater and dried at 80° C. overnight. The template was removed bycalcination in air at 750° C. for 6 h with a ramp up of 4° C. per min.

Catalyst Properties

Table 1 includes mechanical properties of the catalysts prepared inExamples 1 and 2.

TABLE 1 BET Surface Area Pore Volume Catalysts/Supports (m²/g) (cm³/g)Silica-Alumina Catalyst 100% SiO₂ - 0% Al₂O₃ 304.41 1.13 95% SiO₂ - 5%Al₂O₃ 303.72 1.09 90% SiO₂ - 10% Al₂O₃ 305.45 0.94 75% SiO₂ - 25% Al₂O₃282.37 1.00 50% SiO₂ - 50% Al₂O₃ 253.46 0.95 25% SiO₂ - 75% Al₂O₃ 231.980.79 0% SiO₂ - 100% Al₂O₃ 204.82 0.70 Mesoporous Silica CatalystWO₃/100% SiO₂ - 0% 374.35 0.81 Al₂O₃

Catalyst Evaluation

The prepared catalysts from Examples 1-4 were tested for their activityand selectivity to butene in a fixed bed continuous flow reactor (ID0.25 in, Autoclave Engineers Ltd.) at atmospheric pressure. A fixedamount of catalyst samples, 1 ml of each catalyst type (with a total of3 ml) was packed in the reactor tube with silicon carbide at the bottomof the reactor. The silicon carbide is inert and makes no contributionto the reaction chemistry. Each catalyst type and the silicon carbidewere separated by quartz wool with an additional layer of quartz wool atthe top and bottom of the reactor. The catalysts were pretreated andactivated under N₂ at 550° C. and a flow of 25 standard cubiccentimeters per minute (sccm) for 1 hour. All reactions were carried outat three temperature of 450° C., 500° C., and 550° C., a GHSV (gashourly space velocity) of 900 h⁻¹, and atmospheric pressure using2-butene (5 milliliters/minutes (ml/min)) as feed with nitrogen asdiluent (25 ml/min). The temperature was maintained for 3.5 hours. Thequantitative analysis of the reaction products were carried out on-lineusing an Agilent gas chromatograph with flame ionization detector (FID)(Agilent GC-7890B), equipped with an HP-Al/KCL (50 m×0.53 mm×15 microns)column.

Tables 2-4 indicate the catalytic performance of the silica-aluminacatalyst individually. The silica-alumina catalysts were preparedaccording to Example 1. The silica-alumina catalyst was screened forisomerization activity as well as any propylene production contribution.Each of the silica-alumina catalysts with varying silica to aluminaratios were tested at three temperatures of 500° C., 525° C., and 550°C. represented by Tables 2, 3, and 4 respectively. Values of the yieldsand conversions of the 2-butene feed were calculated based on an averageof values obtained from 5 injections into the gas chromatograph at eachtemperature.

TABLE 2 Trans- 1- Iso- Cis- Silica-Alumina Ethylene Propylene ButeneButene Butene Butene C5 C6+ Conversion Ratio (mol %) (mol %) (mol %)(mol %) (mol %) (mol %) (mol %) (mol %) (%) Temperature = 500° C. 100%SiO₂- 0.00 0.00 48.55 9.07 0.25 42.16 0.00 0.00 9.29 0% Al₂O₃ 95% SiO₂-0.12 0.74 37.60 27.20 6.89 27.88 0.25 0.00 34.52 5% Al₂O₃ 90% SiO₂- 0.000.00 35.61 25.53 12.51 26.35 0.00 0.00 38.04 10% Al₂O₃ 75% SiO₂- 0.000.88 35.21 24.81 13.66 26.07 0.14 0.00 38.73 25% Al₂O₃ 50% SiO₂- 0.231.31 32.18 24.31 15.61 24.27 1.68 0.00 43.56 50% Al₂O₃ 25% SiO₂- 0.351.52 32.07 23.23 17.22 23.65 1.68 0.00 44.29 75% Al₂O₃ 0% SiO₂- 0.080.50 37.24 26.65 7.40 27.59 0.60 0.00 35.17 100% Al₂O₃

TABLE 3 Trans- 1- Iso- Cis- Silica-Alumina Ethylene Propylene ButeneButene Butene Butene C5 C6+ Conversion Ratio (mol %) (mol %) (mol %)(mol %) (mol %) (mol %) (mol %) (mol %) (%) Temperature = 525° C. 100%SiO₂- 0.00 0.00 48.42 9.60 0.28 41.70 0.00 0.00 9.88 0% Al₂O₃ 95% SiO₂-0.00 0.23 37.20 28.69 5.59 27.97 0.49 0.00 34.83 5% Al₂O₃ 90% SiO₂- 0.000.00 34.27 26.06 13.43 25.56 0.69 0.00 40.18 10% Al₂O₃ 75% SiO₂- 0.000.00 34.27 25.72 13.24 25.57 1.19 0.00 40.15 25% Al₂O₃ 50% SiO₂- 0.210.95 33.18 26.52 11.58 25.11 1.83 0.00 41.71 50% Al₂O₃ 25% SiO₂- 0.001.07 34.27 26.14 10.72 25.71 2.86 0.00 40.03 75% Al₂O₃ 0% SiO₂- 0.130.48 37.10 28.07 4.40 27.72 1.68 0.00 35.18 100% Al₂O₃

TABLE 4 Trans- 1- Iso- Cis- Silica-Alumina Ethylene Propylene ButeneButene Butene Butene C5 C6+ Conversion Ratio (mol %) (mol %) (mol %)(mol %) (mol %) (mol %) (mol %) (mol %) (%) Temperature = 550° C. 100%SiO₂- 0.00 0.00 48.18 10.43 0.32 41.07 0.00 0.00 10.75 0% Al₂O₃ 95%SiO₂- 0.00 0.41 35.68 29.15 6.49 27.04 0.95 0.00 37.28 5% Al₂O₃ 90%SiO₂- 0.00 0.00 32.68 26.35 14.85 24.61 1.51 0.00 42.70 10% Al₂O₃ 75%SiO₂- 0.00 0.77 32.60 25.90 13.66 24.57 2.51 0.00 42.84 25% Al₂O₃ 50%SiO₂- 0.31 1.22 32.06 27.09 9.80 24.43 3.86 0.00 43.52 50% Al₂O₃ 25%SiO₂- 0.40 1.03 31.95 25.87 9.39 24.17 6.78 0.00 43.88 75% Al₂O₃ 0%SiO₂- 0.14 0.70 35.01 28.38 3.35 26.66 4.60 0.00 38.33 100% Al₂O₃

Tables 2-4 demonstrate that increasing the Al₂O₃ to SiO₂ ratio over 10%Al₂O₃ does not correlate to greater isomerization activity.Specifically, despite increasing the Al₂O₃ content to 25%, 50%, and 75%,the 1-butene yield remains steady at around 26. Further, the amount of1-butene formed at each of the three tested temperatures is notsubstantially different for all seven samples with the exception of the100% SiO₂ catalyst which exhibits a lesser isomerization activity. Asthe Al₂O₃ ratio increases, in general the amount of undesired isobutenealso increases. Overall, it may be determined from Tables 2-4 that adesirable ratio of Al₂O₃ to SiO₂ is 95% SiO₂ and 5% Al₂O₃ as maximum1-butene and minimal isobutene is achieved. As a general rule, thehigher the alumina content of the catalyst, the quicker the deactivationof the catalyst. While it is noted that the 100% Al₂O₃ isomerizes a morethan the catalysts that contain both silica and alumina, the 100% Al₂O₃catalyst is the quickest to deactivate, the tradeoff is not justifiable.Furthermore, the high alumina content catalyst results in cracking andmetathesis at higher temperatures, which is undesirable in theisomerization catalyst zone.

Table 5 indicates the catalytic performance of the mesoporous silicacatalyst of Example 2. The mesoporous silica catalyst was screened formetathesis activity as well as any propylene production contribution.The mesoporous silica catalyst was tested at three temperatures of 500°C., 525° C., and 550° C. Values of the yields and conversions of the2-butene feed were calculated based on an average of values obtainedfrom 5 injections into the gas chromatograph at each temperature.

TABLE 5 Trans- 1- Iso- Cis- Ethylene Propylene Butene Butene ButeneButene C5 C6+ Conversion Temperature (mol %) (mol %) (mol %) (mol %)(mol %) (mol %) (mol %) (mol %) (%) Sample = 10% WO₃/100% SiO₂-0% Al₂O₃500° C. 1.60 18.81 22.73 13.90 0.42 16.77 20.21 5.56 60.50 525° C. 2.5823.53 17.88 10.48 0.29 13.25 22.28 9.70 68.86 550° C. 3.41 25.48 16.0210.71 0.38 12.01 21.59 10.12 71.97

Table 5 demonstrates the increasing trend in the yield of desirablepropylene formation with temperature increases. Conversely, anincreasing trend in the yield of less desirable C₆+ hydrocarbons is alsoobserved with temperature increases. As such, running the reactor at550° C. provides the desirable benefit of an increased propyleneproduction with an acceptable C6+ production given the increasedpropylene output. Running the reactor at 500° C. is believed topotentially provide the benefit of an increased lifetime of the catalystbed associated with a lesser operating temperature. The lesser operatingtemperature in the range of 500° C. still achieves acceptableisomerization activity from the silica-alumina catalyst in conjunctionwith the acceptable metathesis activity from the mesoporous silicacatalyst. In total, operating temperatures of 500° C., 525° C., and 550°C. each exhibit desirable propylene production and demonstrate the widerange of reactor temperatures in which the multiple-stage catalystsystem may operate

Table 6 indicates the catalytic performance of the MFI structured silicacatalyst of Example 4. The tested MFI structured silica catalyst wasZSM-5 with Si/Al=2000 (MFI-2000). The mesoporous silica catalyst wasscreened for cracking activity as well as any propylene productioncontribution. The mesoporous silica catalyst was tested at twotemperatures of 500° C. and 550° C. Values of the yields and conversionsof the 2-butene feed were calculated based on an average of valuesobtained from 5 injections into the gas chromatograph at eachtemperature.

TABLE 6 Trans- 1- Iso- Cis- Ethylene Propylene Butene Butene ButeneButene C5 C6+ Conversion Temperature (mol %) (mol %) (mol %) (mol %)(mol %) (mol %) (mol %) (mol %) (%) Sample = ZSM-5 500° C. 14.343 37.7735.865 3.925 9.867 4.284 6.946 7.401 89.852 550° C. 17.534 41.771 5.5634.135 8.945 4.137 4.023 6.746 90.300

Table 6 demonstrates the increasing trend in the yield of desirablepropylene formation with temperature increases. Similarly, a decreasingtrend in the yield of less desirable C6+ hydrocarbons is also observedwith temperature increases. As such, running the reactor at 550° C.provides the desirable benefit of an increased propylene production.

Tables 2-6 provide each of the isomerization, metathesis, and crackingcatalysts of Examples 1, 2, and 4 respectively tested separately assingle catalyst systems. Specifically, the isomerization, metathesis,and cracking catalysts were individually test without being combinedwith one or more of the other catalyst type. This provided the effect ofthe isomerization, metathesis, and cracking catalysts individually tothe combined multiple-stage catalyst system with all three of theisomerization, metathesis, and cracking catalysts.

As illustrated in Table 7, a comparative dual catalyst systemconfigurations was also tested. The optimal dual catalyst configurationwas a catalyst system consisting of the metathesis catalyst (mesoporoussilica catalyst) on the top and the cracking catalyst (MFI structuredsilica catalyst) in the bottom of the reactor. This configurationrepresents the multiple-stage catalyst system in the absence of theisomerization catalyst (silica-alumina catalyst). Specifically, thecomparative dual catalyst system comprised a mesoporous silica catalystof 10% WO₃/100% SiO₂−0% Al₂O₃ (Example 2) and a MFI structured silicacatalyst of ZSM-5 with a Si/Al ratio equal to 2000 (MFI-2000) (Example4). This optimized comparative dual catalyst system was subsequentlycompared to a triple bed catalyst system in accordance with the presentdisclosure. The triple bed catalyst system was a silica-alumina catalystof 5% Al₂O₃+95% SiO₂ (Example 1), a mesoporous silica catalyst of 10%WO₃/100% SiO₂−0% Al₂O₃ (Example 2), and a MFI structured silica catalystof ZSM-5 with a Si/Al ratio equal to 2000 (MFI-2000) (Example 4).

TABLE 7 Triple Catalyst Comparative Dual Catalyst % Yield Temperature °C. Change Difference 450 500 550 450 500 550 Triple vs. Double Yield(mol. %) Methane 0.000 0.114 0.300 0.000 0.101 0.238 0.06 Ethane 0.0830.128 0.187 0.074 0.102 0.151 0.04 Ethylene 8.243 11.931 16.072 7.10710.391 14.489  11% 1.58 Propane 3.365 2.436 1.727 2.801 2.232 1.727 0.00Propylene 33.408 39.558 44.425 33.409 38.897 43.360 2.5% 1.07 Iso-Butane4.532 2.101 0.889 4.101 2.324 1.131 −0.24 N-Butane 2.629 1.521 0.8802.430 1.710 1.039 −0.16 Trans-Butene 6.972 6.632 6.149 7.670 2.340 7.073−0.92 1-Butene 4.414 4.763 4.957 4.117 4.402 4.579 0.38 Iso-Butene12.646 11.382 10.135 13.220 12.013 10.960  −8% −0.82 Cis-Butene 5.0404.887 4.620 5.371 5.250 5.119 −0.50 C₅ 10.320 7.393 4.165 13.100 10.1326.718 −2.55 C₆+ 8.366 7.153 5.493 6.537 5.126 3.418 2.08 Total Olefins41.651 51.489 60.493 40.516 49.288 57.849   5% 2.65 (C₃═ & C₂═)Conversion (mol. %) Conversion 87.987 88.481 89.231 86.959 87.410 87.8081.42 Conversion-C₄ 70.927 72.335 74.139 69.522 70.994 72.270 1.87Selectivity Propylene 37.986 44.709 49.787 38.420 44.301 49.379   1%0.41 Selectivity Ethylene 9.365 13.484 18.012 8.172 11.888 16.499 9.2%1.51 Selectivity Isobutene 14.373 12.864 11.358 15.203 13.744 12.481 −9% −1.12 Selectivity

Table 7 demonstrates the superior performance of the triple bed catalystsystem in accordance with the present disclosure over a comparative dualcatalyst system. The triple bed catalyst system comprised the dualcatalyst system (Examples 2 and 4) with the added isomerization catalystzone of a silica-alumina catalyst of 5% Al₂O₃+95% SiO₂ (Example 1).Compared to the dual catalyst system, introduction of the isomerizationlayer in the triple catalyst system on top of the metathesis andcracking catalyst layers resulted in both the desirable ethylene andpropylene yields increasing and the undesirable isobutene yielddecreasing.

With reference to FIG. 3, the multiple-stage catalyst systemdemonstrates a stable conversion and selectivity at various temperaturesover a period of time. A triple catalyst system in accordance with thepresent disclosure (a silica-alumina catalyst of 5% Al₂O₃+95% SiO₂(Example 1), a mesoporous silica catalyst of 10% WO₃/100% SiO₂−0% Al₂O₃(Example 2), and a MFI structured silica catalyst of ZSM-5 with a Si/A1ratio equal to 2000 (MFI-2000) (Example 4)) was utilized for thereaction of 2-butene to propylene at reaction temperatures of 450, 500and 550° C. at atmospheric pressure with a GHSV of 900 h⁻¹. The data ofFIG. 3 is also presented as Table 8.

TABLE 8 Temperature = 450° C. Time on Stream 0:22:14 1:08:28 1:54:442:41:00 3:27:17 Conversion 88.39 88.10 88.05 87.70 87.71 Propylene 39.3537.36 37.37 38.17 37.59 Selectivity Ethylene 11.73 9.63 8.73 8.54 8.19Selectivity IsoButene 10.95 11.01 11.01 11.33 11.33 SelectivityTemperature = 500° C. Time on Stream 4:13:32 4:59:48 5:46:02 6:32:147:18:30 Conversion 88.37 88.63 88.61 88.2 88.11 Propylene 42.47 44.0944.61 45.43 45.64 Selectivity Ethylene 11.84 13.47 13.49 13.48 13.34Selectivity IsoButene 10.39 9.97 9.95 10.30 10.39 SelectivityTemperature = 550° C. Time on Stream 8:04:47 8:51:03 9:37:18 10:23:3511:09:51 Conversion 88.78 89.37 89.33 89.26 89.20 Propylene 47.98 49.1549.66 49.73 49.84 Selectivity Ethylene 16.03 18.13 18.18 18.09 17.93Selectivity IsoButene 9.59 8.98 8.97 9.02 9.07 Selectivity Temperature =450° C. (Return) Time on Stream 11:56:07 12:42:23 13:28:39 14:14:54Conversion 88.28 86.50 86.54 86.43 Propylene 47.45 39.69 39.44 39.15Selectivity Ethylene 14.86 7.15 6.81 6.61 Selectivity IsoButene 10.1012.21 12.24 12.35 Selectivity

Table 9 indicates the catalytic performance of a triple bed catalystsystem in accordance with this disclosure. Specifically, asilica-alumina catalyst of 5% Al₂O₃+95% SiO₂ (Example 1), a mesoporoussilica catalyst of 10% WO₃/100% SiO₂−0% Al₂O₃ (Example 2), and a MFIstructured silica catalyst of ZSM-5 with a Si/Al ratio equal to 2000(MFI-2000) (Example 4) were utilized in combination. The effect ofvarying the amount of the isomerization catalyst (Example 1) wasdetermined. The triple catalyst system was run with 1 ml of eachcatalyst in the triple catalyst system and was also run with 2 ml of thesilica-alumina catalyst and 1 ml of each of the MFI structured silicacatalyst and mesoporous silica catalyst. The comparative test determinedthat the overall catalyst performance remains substantially unchangedwhen the isomerization catalyst is increased by a factor of 2. Theslight decrease in propylene yield is attributed to the resulting changein the hourly space velocity from the increased isomerization catalyst.

TABLE 9 1 ml Silica-Alumina Catalyst + 1 ml 2 ml Silica-AluminaCatalyst + 1 ml Mesoporous Silica Catalyst + 1 ml MFI Mesoporous SilicaCatalyst + 1 ml MFI Structured Catalyst Structured Catalyst Temperature° C. 450 500 550 450 500 550 Yield (mol. %) Methane 0.000 0.114 0.3000.000 0.114 0.294 Ethane 0.083 0.128 0.187 0.083 0.127 0.188 Ethylene8.243 11.931 16.072 8.233 11.731 15.991 Propane 3.365 2.436 1.727 3.3732.393 1.828 Propylene 33.408 39.558 44.425 32.853 39.270 43.810Iso-Butane 4.531 2.201 0.889 4.547 2.118 1.022 N-Butane 2.629 1.5210.880 2.582 1.551 0.997 Trans-Butene 6.972 6.632 6.149 7.109 6.792 6.2971-Butene 4.414 4.763 4.957 4.180 4.565 4.775 Iso-Butene 12.646 11.38210.135 12.695 11.457 10.204 Cis-Butene 5.040 4.887 4.620 5.080 4.9414.683 C₅ 10.320 7.393 4.165 10.702 7.659 4.405 C₆+ 8.366 7.153 5.4938.580 7.283 5.507 Total Olefins 41.651 51.489 60.498 41.087 51.00159.801 (C₃═ & C₂═) Conversion (mol. %) Conversion 87.987 88.481 89.23187.812 88.267 89.020

Calculation Methodologies

Determination of “% Change” was calculated according to formula 1.

$\begin{matrix}{\frac{{{{yield}@550}\mspace{14mu} C\mspace{14mu} {for}\mspace{14mu} {triple}\mspace{14mu} {bed}} - {{{yield}@550}\mspace{14mu} C\mspace{14mu} {for}\mspace{14mu} {dual}\mspace{14mu} {bed}}}{{{yield}@550}\mspace{14mu} C\mspace{14mu} {for}\mspace{14mu} {dual}\mspace{14mu} {bed}} \times 100} & (1)\end{matrix}$

Determination of “Conversion” was calculated according to formula 2,where n_(i) is the number of moles of component “i” entering or leavingthe reactor.

$\begin{matrix}{{Conversion} = {\frac{n_{i,{i\; n}} - n_{i,{out}}}{n_{i,{i\; n}}} \times 100}} & (2)\end{matrix}$

Similarly, determination of “Conversion-C4” was calculated according toformula 3.

Conversion-C4=100−(CisButene Yield+TransButene Yield+IsoButeneYield+1-Butene Yield)   (3)

Determination of “Selectivity” was calculated according to formula 4.

$\begin{matrix}{{Selectivity} = {\frac{{Yield}\mspace{14mu} {of}\mspace{14mu} {Product}}{Conversion} \times 100}} & (4)\end{matrix}$

The surface area of the samples was measured by nitrogen adsorption at77 K using AUTOSORB-1 (Quanta Chrome). Before adsorption measurements,samples (ca. 0.1 g) were heated at 220° C. for 2 hours under nitrogenflow. The nitrogen adsorption isotherms of catalysts were measured atliquid nitrogen temperature (77 K). The surface areas was calculated bythe Brunauer Emmett-Teller (BET) method. The total pore volume wasestimated from the amount of N₂ adsorbed at P/P0=0.99. Barret E P,Joyner L J, Halenda P H, J. Am. Chem. Soc. 73 (1951) 373-380.

In a first aspect, the disclosure provides a process for production ofpropylene comprising introducing a hydrocarbon stream comprising2-butene to an isomerization catalyst zone to isomerize the 2-butene to1-butene, where the isomerization catalyst zone comprises asilica-alumina catalyst with a ratio by weight of alumina to silica from1:99 to 20:80; passing the 2-butene and 1-butene to a metathesiscatalyst zone to cross-metathesize the 2-butene and 1-butene into ametathesis product stream comprising propylene and C₄-C₆ olefins, wherethe metathesis catalyst comprises a mesoporous silica catalyst supportimpregnated with metal oxide; and cracking the metathesis product streamin a catalyst cracking zone to produce propylene, where the catalystcracking zone comprises a mordenite framework inverted (MFI) structuredsilica catalyst.

In a second aspect, the disclosure provides a process of the firstaspect, in which the silica-alumina catalyst includes a surface area of200 m²/g to 600 m²/g.

In a third aspect, the disclosure provides a process of either the firstor second aspects, in which the silica-alumina catalyst includes asurface area of 250 m²/g to 325 m²/g.

In a fourth aspect, the disclosure provides a process of any one of thefirst through third aspects, in which the silica-alumina catalyst has apore volume of at least 0.60 cm³/g.

In a fifth aspect, the disclosure provides a process of any one of thefirst through fourth aspects, in which the silica-alumina catalyst has apore volume of 0.90 cm³/g to 1.2 cm³/g.

In a sixth aspect, the disclosure provides a process of any one of thefirst through fifth aspects, in which the silica-alumina catalystcomprises an alumina to silica weight ratio between 1:99 and 10:90.

In a seventh aspect, the disclosure provides a process of any one of thefirst through sixth aspects, in which the silica-alumina catalystcomprises an alumina to silica weight ratio between 3:97 and 7:93.

In an eighth aspect, the disclosure provides a process of any one of thefirst through seventh aspects, in which the metal oxide of themesoporous silica catalyst comprises one or more oxides of molybdenum,rhenium, tungsten, or combinations thereof.

In a ninth aspect, the disclosure provides a process of any one of thefirst through eighth aspects, in which the metal oxide of the mesoporoussilica catalyst is tungsten oxide (WO₃).

In a tenth aspect, the disclosure provides a process of any one of thefirst through ninth aspects, in which the mesoporous silica catalystcomprises 1 to 30 weight percent tungsten oxide.

In an eleventh aspect, the disclosure provides a process of any one ofthe first through tenth aspects, in which the mesoporous silica catalystcomprises 8 to 12 weight percent tungsten oxide.

In a twelfth aspect, the disclosure provides a process of any one of thefirst through eleventh aspects, in which the mesoporous silica catalystincludes a surface area of 200 m²/g to 600 m²/g.

In a thirteenth aspect, the disclosure provides a process of any one ofthe first through twelfth aspects, in which the mesoporous silica has apore volume of at least 0.60 cm³/g.

In a fourteenth aspect, the disclosure provides a process of any one ofthe first through thirteenth aspects, in which the MFI structured silicacatalyst includes a total acidity of 0.001 mmol/g to 0.1 mmol/g.

In a fifteenth aspect, the disclosure provides a process of any one ofthe first through fourteenth aspects, in which the MFI structured silicacatalyst comprises alumina.

In a sixteenth aspect, the disclosure provides a process of any one ofthe first through fourteenth aspects, in which the MFI structured silicacatalyst is alumina free.

In a seventeenth aspect, the disclosure provides a process of any one ofthe first through fifteenth aspects, in which the MFI structured silicacatalyst comprises ZSM-5.

In an eighteenth aspect, the disclosure provides a process of any one ofthe first through seventeenth aspects, in which the hydrocarbon feed isa Raffinate 2 stream from a fluid catalytic cracker or an ethylenecracker.

In a nineteenth aspect, the disclosure provides a process of any one ofthe first through seventeenth aspects, in which the hydrocarbon feed isa Raffinate 2 stream which consists of 45 to 55 wt % 1-butene, 20 to 30wt % 2-butene, 10 to 20 wt % n-butane, 5 to 15 wt % iso-butane, and 0 to5 wt % other components.

In a twentieth aspect, the disclosure provides a process of any one ofthe first through seventeenth aspects, in which the hydrocarbon feed isa Raffinate 2 stream which consists of 10 to 20 wt % 1-butene, 20 to 30wt % 2-butene, 8 to 18 wt % n-butane, 37 to 47 wt % iso-butane, and 0 to8 wt % other components.

In an twenty-first aspect, the disclosure provides a multiple-stagecatalyst system for producing propylene from a hydrocarbon stream, themultiple-stage catalyst system comprising an isomerization catalystzone, a metathesis catalyst zone downstream of the isomerization zone,and a cracking catalyst zone downstream of the metathesis catalyst zone;where the isomerization catalyst zone comprises a silica-aluminacatalyst with a ratio by weight of alumina to silica from 1:99 to 20:80,where the silica-alumina catalyst zone isomerizes the 2-butene to1-butene; the metathesis catalyst zone comprises a mesoporous silicacatalyst support impregnated with metal oxide to form a mesoporoussilica catalyst, where the mesoporous silica catalyst zonecross-metathesizes the 2-butene and 1-butene into a metathesis productstream comprising propylene and C₄-C₆ olefins; and the cracking catalystzone comprises a mordenite framework inverted (MFI) structured silicacatalyst, where the cracking catalyst zone cracks the metathesis productstream to produce propylene.

In a twenty-second aspect, the disclosure provides a multiple-stagecatalyst system of the twenty-first aspect, in which the isomerizationcatalyst zone, the metathesis catalyst zone, and the cracking catalystzone are disposed in one reactor.

In a twenty-third aspect, the disclosure provides a multiple-stagecatalyst system of the twenty-first aspect, in which the isomerizationcatalyst zone is disposed on a first reactor, the metathesis catalystzone is disposed in a second reactor downstream of the first reactor,and the cracking catalyst zone is disposed in a third reactor downstreamof the second reactor.

In a twenty-fourth aspect, the disclosure provides a multiple-stagecatalyst system of the twenty-first aspect, in which the isomerizationcatalyst zone is disposed on a first reactor, the cracking catalyst zoneis disposed in a second reactor downstream of the first reactor, and themetathesis catalyst zone is disposed in the first reactor downstream ofthe isomerization catalyst zone or in the second reactor upstream of thecracking catalyst zone.

In a twenty-fifth aspect, the disclosure provides a multiple-stagecatalyst system of any one of the twenty-first through twenty-fourthaspects, in which the silica-alumina catalyst includes a surface area of200 m²/g to 600 m²/g.

In a twenty-sixth aspect, the disclosure provides a multiple-stagecatalyst system of any one of the twenty-first through twenty-fifthaspects, in which the silica-alumina catalyst includes a surface area of250 m²/g to 325 m²/g.

In a twenty-seventh aspect, the disclosure provides a multiple-stagecatalyst system of any one of the twenty-first through twenty-fifthaspects, in which the silica-alumina catalyst has a pore volume of atleast 0.60 cm²/g.

In a twenty-eighth aspect, the disclosure provides a multiple-stagecatalyst system of any one of the twenty-first through twenty-sixthaspects, in which the silica-alumina catalyst has a pore volume of 0.90cm³/g to 1.2 cm³/g.

In a twenty-ninth aspect, the disclosure provides a multiple-stagecatalyst system of any one of the twenty-first through twenty-eighthaspects, in which the silica-alumina catalyst comprises an alumina tosilica weight ratio between 1:99 and 10:90.

In a thirtieth aspect, the disclosure provides a multiple-stage catalystsystem of any one of the twenty-first through twenty-ninth aspects, inwhich the silica-alumina catalyst comprises an alumina to silica weightratio between 3:97 and 7:93.

In a thirty-first aspect, the disclosure provides a multiple-stagecatalyst system of any one of the twenty-first through thirtiethaspects, in which where the metal oxide of the mesoporous silicacatalyst comprises one or more oxides of molybdenum, rhenium, tungsten,or combinations thereof.

In a thirty-second aspect, the disclosure provides a multiple-stagecatalyst system of any one of the twenty-first through thirty-firstaspects, in which the metal oxide of the mesoporous silica catalyst istungsten oxide (WO₃).

In a thirty-third aspect, the disclosure provides a multiple-stagecatalyst system of any one of the twenty-first through thirty-second, inwhich the mesoporous silica catalyst comprises 8 to 12 weight percenttungsten oxide.

In a thirty-fourth aspect, the disclosure provides a multiple-stagecatalyst system of any one of the twenty-first through thirty-thirdaspects, in which the mesoporous silica catalyst includes a surface areaof 200 m²/g to 600 m²/g.

In a thirty-fifth aspect, the disclosure provides a multiple-stagecatalyst system of any one of the twenty-first through thirty-fourthaspects, in which the mesoporous silica catalyst includes a surface areaof 250 m²/g to 300 m²/g.

In a thirty-sixth aspect, the disclosure provides a multiple-stagecatalyst system of any one of the twenty-first through thirty-fifthaspects, in which the mesoporous silica catalyst has a pore volume of atleast 0.60 cm³/g.

In a thirty-seventh aspect, the disclosure provides a multiple-stagecatalyst system of any one of the twenty-first through thirty-sixthaspects, in which the mesoporous silica catalyst has a pore volume of0.80 cm³/g to 1.3 cm³/g.

In a thirty-eighth aspect, the disclosure provides a multiple-stagecatalyst system of any one of the twenty-first through thirty-seventhaspects, in which the MFI structured silica catalyst includes a totalacidity of 0.001 mmol/g to 0.1 mmol/g.

In a thirty-ninth aspect, the disclosure provides a multiple-stagecatalyst system of any one of the twenty-first through thirty-eighthaspects, in which the MFI structured silica catalyst comprises alumina.

In a fortieth aspect, the disclosure provides a multiple-stage catalystsystem of any one of the twenty-first through thirty-ninth aspects, inwhich the MFI structured silica catalyst is alumina free.

In a forty-first aspect, the disclosure provides a multiple-stagecatalyst system of any one of the twenty-first through thirty-ninthaspects, in which the MFI structured silica catalyst comprises ZSM-5.

Throughout this disclosure ranges are provided for various parametersand characteristics of the catalysts and multiple-stage catalyst system.It will be appreciated that when one or more explicit ranges areprovided the individual values and the ranges formed therebetween arealso intended to be provided as providing an explicit listing of allpossible combinations is prohibitive. For example, a provided range of1-10 also includes the individual values, such as 1, 2, 3, 4.2, and 6.8,as well as all the ranges which may be formed within the providedbounds, such as 1-8, 2-4, 6-9, and 1.3-5.6.

It should now be understood that various aspects of the systems andmethods of making propylene with the multiple catalysts are describedand such aspects may be utilized in conjunction with various otheraspects. It should also be understood to those skilled in the art thatvarious modifications and variations can be made to the describedembodiments without departing from the spirit and scope of the claimedsubject matter. Thus, it is intended that the specification cover themodifications and variations of the various described embodimentsprovided such modification and variations come within the scope of theappended claims and their equivalents.

What is claimed is:
 1. A process for production of propylene comprising:introducing a hydrocarbon stream comprising 2-butene to an isomerizationcatalyst zone to isomerize the 2-butene to 1-butene, where theisomerization catalyst zone comprises a silica-alumina catalyst with aratio by weight of alumina to silica from 1:99 to 20:80; passing the2-butene and 1-butene to a metathesis catalyst zone to cross-metathesizethe 2-butene and 1-butene into a metathesis product stream comprisingpropylene and C₄-C₆ olefins, where the metathesis catalyst comprises amesoporous silica catalyst support impregnated with metal oxide; andcracking the metathesis product stream in a catalyst cracking zone toproduce propylene, where the catalyst cracking zone comprises amordenite framework inverted (MFI) structured silica catalyst.
 2. Theprocess of claim 1 where the silica-alumina catalyst includes a surfacearea of 200 m²/g to 600 m²/g.
 3. The process of claim 1 where thesilica-alumina catalyst has a pore volume of at least 0.60 cm³/g.
 4. Theprocess of claim 1 where the silica-alumina catalyst comprises analumina to silica weight ratio between 1:99 and 10:90.
 5. The process ofclaim 1 where the metal oxide of the mesoporous silica catalystcomprises one or more oxides of molybdenum, rhenium, tungsten, orcombinations thereof.
 6. The process of claim 1 where the metal oxide ofthe mesoporous silica catalyst is tungsten oxide (WO₃).
 7. The processof claim 1 where the mesoporous silica has a pore volume of at least0.60 cm³/g.
 8. The process of claim 1 where the MFI structured silicacatalyst includes a total acidity of 0.001 mmol/g to 0.3 mmol/g.
 9. Theprocess of claim 1 where the hydrocarbon feed is a Raffinate 2 streamfrom a fluid catalytic cracker or an ethylene cracker.
 10. Amultiple-stage catalyst system for producing propylene from ahydrocarbon stream comprising 2-butene, the multiple-stage catalystsystem comprising an isomerization catalyst zone, a metathesis catalystzone downstream of the isomerization zone, and a cracking catalyst zonedownstream of the metathesis catalyst zone where: the isomerizationcatalyst zone comprises a silica-alumina catalyst with a ratio by weightof alumina to silica from 1:99 to 20:80, where the silica-aluminacatalyst zone isomerizes the 2-butene to 1-butene; the metathesiscatalyst zone comprises a mesoporous silica catalyst support impregnatedwith metal oxide to form a mesoporous silica catalyst, where themesoporous silica catalyst zone cross-metathesizes the 2-butene and1-butene into a metathesis product stream comprising propylene and C₄-C₆olefins; and the cracking catalyst zone comprises a mordenite frameworkinverted (MFI) structured silica catalyst, where the cracking catalystzone cracks the metathesis product stream to produce propylene.
 11. Themultiple-stage catalyst system of claim 10 where the isomerizationcatalyst zone, the metathesis catalyst zone, and the cracking catalystzone are disposed in one reactor.
 12. The multiple-stage catalyst systemof claim 10 where the isomerization catalyst zone is disposed on a firstreactor, the metathesis catalyst zone is disposed in a second reactordownstream of the first reactor, and the cracking catalyst zone isdisposed in a third reactor downstream of the second reactor.
 13. Themultiple-stage catalyst system of claim 10 where the isomerizationcatalyst zone is disposed on a first reactor, the cracking catalyst zoneis disposed in a second reactor downstream of the first reactor, and themetathesis catalyst zone is disposed in the first reactor downstream ofthe isomerization catalyst zone or in the second reactor upstream of thecracking catalyst zone.
 14. The multiple-stage catalyst system of claim10 where the silica-alumina catalyst includes a surface area of 200 m²/gto 600 m²/g.
 15. The multiple-stage catalyst system of claim 10 wherethe silica-alumina catalyst has a pore volume of at least 0.60 cm²/g.16. The multiple-stage catalyst system of claim 10 where thesilica-alumina catalyst comprises an alumina to silica weight ratiobetween 1:99 and 10:90.
 17. The multiple-stage catalyst system of claim10 where the metal oxide of the mesoporous silica catalyst comprises oneor more oxides of molybdenum, rhenium, tungsten, or combinationsthereof.
 18. The multiple-stage catalyst system of claim 10 where themetal oxide of the mesoporous silica catalyst is tungsten oxide (WO₃).19. The multiple-stage catalyst system of claim 10 where the mesoporoussilica catalyst has a pore volume of at least 0.60 cm³/g.
 20. Themultiple-stage catalyst system of claim 10 where the MFI structuredsilica catalyst includes a total acidity of 0.001 mmol/g to 0.3 mmol/g.